Solvent refined coal process with retention of coal minerals

ABSTRACT

A solvation process for producing deashed solid and liquid hydrocarbonaceous fuel from coal. Raw coal is slurred with a solvent comprising hydroaromatic compounds in contact with hydrogen in a first zone to dissolve hydrocarbonaceous fuel from coal minerals by transfer of hydrogen from hydroaromatic solvent compounds to hydrocarbonaceous material in the coal. The slurry is then treated with hydrogen in a second zone to replenish the solvent with hydrogen. The process is improved by retention of coal minerals in the second zone.

United States Patent Hinderliter et al.

[ SOLVENT REFINED COAL PROCESS WITH RETENTION 0F COAL MINERALS [75] inventors: Charles R. Hinderliter, Overland Park; Russell E. Perrussel, Merriam, both of Kans.

[73] Assignee: The United States of America as represented by the Secretary oi the Interior, Washington, DC.

[22] Filed: Mar. 4, 1974 [211 App]. No.: 446,974

[52] US. Cl. 208/8 [51] Int. Cl Cl0g 1/00 [58] Field of Search 208/8 [56] References Cited UNITED STATES PATENTS 3,34] ,447 9/1967 Bullet al. 208/8 3,594,304 7/197] Seitzer et al. 208/8 [451 May 20, 1975 3,617,465 ll/l97l Wolk et al. 208/8 3,645,885 2/l972 Harris et al 3,808,] l9 4/1974 Bull et al. 208/8 Primary ExaminerDelbert E. Gantz Assistant Examiner-James W. Hellwege [57] ABSTRACT A solvation process for producing deashed solid and liquid hydrocarbonaceous fuel from coal. Raw coal is slurred with a solvent comprising hydroaromatic compounds in contact with hydrogen in a first zone to dissolve hydrocarbonaceous fuel from coal minerals by transfer of hydrogen from hydroaromatic solvent compounds to hydrocarbonaceous material in the coal. The slurry is then treated with hydrogen in a second zone to replenish the solvent with hydrogen. The process is improved by retention of coal minerals in the second zone.

18 Claims, 3 Drawing Figures PATENTED HAY20975 EFFECT OF MINERALS RECYCLE ON SOLVENT HYDROGEN LEVEL IRON IN FEED SLURRY- PERCENT FIG. 2

EFFECT OF MINERALS RECYCLE AND DISSOLVER TEMPERATURE ON HYDROAROMATIC HYDROGEN ACTIVITY RON 1N FEED SLURRY PERCENT SOLVENT REFINED COAL PROCESS WITH RETENTION OF COAL MINERALS This invention resulted from work performed under Contract No. 14-01-0001-496 between The Pittsburg & Midway Coal Mining Co., a subsidiary of Gulf Oil Corporation, and the Office of Coal Research in the Department of the Interior entered into persuant to the Coal Research Act, 30 USC 661 to 668.

This invention relates to a liquid solvent dissolving process for producing reduced or low ash hydrocarbonaceous solid fuel and hydrocarbonaceous distillate liquid fuel, from ash-containing raw coal. Preferred coal feeds contain hydrogen, such as bituminous and subbituminous coals, and lignites. The process produces deashed solid fuel (dissolved coal) together with as much coal derived liquid fuel as possible, with an increase in liquid fuel product being accompanied by a decrease in solid fuel product. Liquid fuel is the more valuable product but the production of liquid fuel is limited because it is accompanied by production of undesired by-product hydrocarbon gases. Although liquid fuel is of greater economic value than deashed solid fuel, hydrocarbon gases are of smaller economic value than either deashed solid fuel or liquid fuel and have a greater hydrogen to carbon ratio than either solid or liquid fuel so that their production is not only wasteful of other fuel product but is also wasteful of hydrogen.

Hydrocarbon gases are produced primarily by hydrocracking, and since their production is undesired in this process no external catalyst is employed, since catalysts generally impart hydrocracking activity in a coal solvation process.

When raw coal is subjected to solvation at a relatively low temperature, the dissolved product comprises in major proportion a high molecular weight fuel which is solid at room temperature. When the mixture of solvent and dissolved coal is subsequently filtered to remove ash and undissolved coal and the filtrate is then subjected to vacuum distillation, this high boiling solid fuel product is recovered as the vacuum bottoms. This deashed vacuum bottoms product is referred to herein as either vacuum bottoms or deashed solid fuel product. This vacuum bottoms is cooled to room temperature on a conveyor belt and is scraped from the belt as fragmented deashed hydrocarbonaceous solid fuel.

As the temperaturue of the solvation process is progressively increased, the vacuum bottoms (deashed solid fuel), which is a high molecular weight polymer, is converted to lower molecular weight hydrocarbonaceous liquid fuel which is chemically similar to the process solvent and which has a similar boiling range. The liquid fuel product is in part recycled as process solvent for the subsequent pass and is referred to herein as either liquid fuel product or excess solvent. Production of liquid fuel occurs by depolymerization of solid fuel through various reactions, such as removal therefrom of heteroatoms, including sulfur and oxygen. As a result of the depolymerization reactions, the liquid fuel has a somewhat higher hydrogen to carbon ratio than the solid fuel and therefore exhibits a correspondingly higher heat content upon combustion. it is desirable in the process to convert as much of the vacuum bottoms (solid fuel) product to solvent boiling range (liquid fuel) product, since liquid fuel is economically more valuable than solid fuel. As the temperature of solvation continues to be increased, an increasing proportion of vacuum bottoms fuel is converted to solvent boiling range fuel until a temperature is reached at which conversion of vacuum bottoms to liquid fuel occurs only at the price of excessive and wasteful production of relatively hydrogen-rich by-product hydrocarbon gases due to the onset of excessive thermal hydrocracking. The present process produces 20 or 40 to weight percent of deashed solid fuel on an MAF (moisture and ash free) basis, the remaining product being primarily liquid fuel.

it is the purpose of the present invention to avoid thermal hydrocracking as much as possible and at least to the extent of avoiding excessive production of hydrocarbon gases since production of gases diminishes the yield of desired deashed solid fuel and liquid fuel products. This purpose is accomplished according to the present invention by performing the solvation process in two separate stages, each stage preferably employing a different temperature. In one embodiment of this invention, less than 6 weight percent of hydrocarbon gases, based on MAF coal feed, is produced. The production limit of hydrocarbon gases establishes the production limit of liquid product and therefore also the production limit of solid fuel product.

An important advantage of the dual temperature method of this invention is that a high temperature stage is made possible whereby product sulfur level can be reduced. Relatively high temperatures are required for sulfur removal whereas temperatures below the required level are not as effective for sulfur removal. The high temperatures required for effective sulfur removal also induce hydrocracking but the hydrocracking reaction is more time dependent and by rapid reduction of the high process temperatures reduction of sulfur level is achieved with a minimum of hydrocracking.

The first reactor stage of the present process is a tubular preheater having a relatively short residence time in which a slurry of feed coal and solvent in essentially plug flow is progressively increased in temperature as it flows through the tube. The tubular preheater has a length to diameter ratio of at least 100, generally, and at least 1,000, preferably. A series of different reactions occur within a flowing stream increment as the temperature of the increment increases from a low inlet temperature to a maximum or exit temperature, at which it remains for only a short time. The second reactor stage employs a relatively longer residence time in a larger vessel maintained at a substantially uniform temperature throughout. Preferably, a regulated amount of forced cooling occurs between the stages so that the second stage temperaturee is lower than the maximum preheater temperature. Although the preheater stage is operated with plug flow without significant back-mixing, full solution mixing with a uniform reactor temperature occurs in the dissolver stage. Data presented below show that a split temperature coal dissolving process results in high conversion of raw coal to deashed solid fuel and liquid fuel and the proportion of liquid to solid fuel product is enhanced while avoiding excessive production of by-product hydrocarbon gases. It is shown below that these results are better accomplished by employing a split temperature process than by employing a process having a uniform temperature in two stages, even when the uniform temperature is the same as either temperature of a split temperature operation.

The coal solvent for the present process comprises liquid hydroaromatic compounds. The coal is slurried with the solvent for charging to the first or preheater stage. 1n the first stage, hydrogen transfer from the solvent hydroaromatic compounds to coal hydrocarbonaceous material occurs resulting in swelling of the coal and in breaking away of hydrocarbon polymers from coal minerals. The range of maximum temperatures suitable in the first (preheater) stage is generally 400 to 525C, or preferably 425 to 500C. If there are inadequate facilities to handle hydrocarbon gaseous by-product the upper temperature limit should be 470C, or below, in order to minimize production of gaseous product. The residence time in the preheater stage is generally 0.01 to 0.25 hours, or preferably 0.01 to 0.15 hours.

In the second (dissolver) stage of the process of this invention, the solvent compounds, which have been depleted of hydrogen and converted to their precursor aromatics by hydrogen donation to the coal in the first stage, are reacted with gaseous hydrogen and reconverted to hydroaromatics for recycle to the first stage. The temperature in the dissolver stage is 350 to 475C, generally, and 400 to 450C, preferably. The residence time in the dissolver stage is 0.1 to 3.0 hours, generally, and 0.15 to 1.0 hours, preferably. The temperature in the dissolver stage is advantageously lower than the maximum temperature in the preheater stage. Any suitable forced cooling step can be employed to reduce stream temperature between the preheater and the dissolver. For example, makeup hydrogen can be charged to the process between the preheater and dissolver steps or a heat exchanger can be employed. Also, the residence time in the preheater is lower than the residence time in the dissolver.

The liquid space velocity for the process (volume of slurry per hour per volume of reactor) ranges from 0.2 to 8.0, generally, and 0.5 to 3.0, preferably. The ratio of hydrogen to slurry ranges from 200 to 10,000 standard cubic feet per barrel, generally, and 500 to 5,000 standard cubic feet per barrel, preferably, (3.6 to 180 generally, and 9 to 90, preferably SCM/IOOL). The weight ratio of recycled solvent product to coal in the feed slurry ranges from 0.511 to :1, generally, and from 1.021 to 2.5:], preferably.

The reactions in both stages in contact with gaseous hydrogen and in both stages heteroatom sulfur and oxygen are removed from solvated deashed coal polymer, resulting in depolymerization and conversion of dissolved coal polymers to desulfurized and deoxygenated free radicals of reduced molecular weight. These free radicals have a tendency to repolymerize at the high temperatures reached in the preheater stage, but at the reduced temperature of the dissolver stage of this invention these free radicals tend to be stabilized against repolymerization by accepting hydrogen at the free radical site. The present process can employ carbon monoxide and steam together with or in place of hydrogen since carbon monoxide and steam react to form hydrogen. The steam can be derived from feeding wet coal or can be injected as water. The reaction of hydrogen at the free radical site occurs more readily at the relatively low dissolver temperature than at the higher preheater exit temperature.

The solvent used at process start up is advantageously derived from coal. lts composition will vary, depending on the properties of the coal from which it is derived. In general, the solvent is a highly aromatic liquid obtained from previous processing of fuel, and generally boils within the range of about C. to 450C. Other generalized characteristics include a density of about 1.1 and a carbon to hydrogen mole ratio in the range from about 1.0 to 0.9 to about 1.0 to 0.3. Any organic solvent for coal can be used as the start-up solvent in the process. A solvent found particularly useful as a start-up solvent is anthracene oil or creosote oil having a boiling range of about 220C. to 400C. However, the start-up solvent is only a temporary process component since during the process dissolved fractions of the raw coal constitute additional solvent which, when added to start-up solvent, provides a total amount of solvent exceeding the amount of start-up solvent. Thus, the original solvent gradually loses its identity and approaches the constitution of the solvent formed by solution and depolymen'zation of the coal in the process. Therefore, in each pass of the process after startup, the solvent can be considered to be a portion of the liquid product produced in previous extraction of the raw coal.

The residence time for the dissolving step in the preheater stage is critical in the process of this invention. Although the duration of the solvation process can vary for each particular coal treated, viscosity changes as the slurry flows along the length of the preheater tube provide a parameter to define slurry residence time in the preheater stage. The viscosity of an increment of feed solution flowing through the preheater initially increases with increasing increment time in the preheater, followed by a decrease in viscosity as the solubilizing of the slurry is continued. The viscosity would rise again at the preheater temperature, but preheater residence time is terminated before a second relatively large increase in viscosity is permitted to occur. An advantageous means for establishing proper time for completion of the preheater step is use of the Relative Viscosity" of the solution formed in the preheater, which is the ratio of the viscosity of the solution formed to the viscosity of the solvent, as fed to the process, both viscosities being measured at 99C. Accordingly, the term Relative Viscosity as used herein is defined as the viscosity at 99C., of an increment of solution, divided by the viscosity of the solvent alone fed to the system measured at 99C., i.e.

Relative Viscosity Viscosity of Solution at 99C./Viscosity of Solvent at 99C.

The Relative Viscosity" can be employed as an indication of the residence time for the solution in the preheater. As the solubilizing of an increment of slurry proceeds during flow through the preheater, the Relative Viscosity of the solution first rises above a value of 20 to a point at which the solution is extremely viscous and in a gel-like condition. In fact. if low solvent to coal ratios are used, for example, 0.5:1, the slurry would set up into a gel. After reaching the maximum Relative Viscosity", well above the value of 20, the Relative Viscosity of the increment begins to decrease to a minimum, after which it has a tendency to again rise to higher values. The solubilization proceeds until the decrease in Relative Viscosity" (following the initial rise in Relative Viscosity) falls to a value at least below 10, whereupon the preheater residence time is terminated and the solution is cooled and passed to the dissolver stage which is maintained at a lower temperature to prevent the Relative Viscosity" from again rising above 10. Normally, the decrease in Relative Viscosity will be allowed to proceed to a value less than 5 and preferably to the range of 1.5 to 2. The conditions in the preheater are such that Relative Vis cosity will again increase to a value about 10, absent abrupt termination of preheater exit conditions, such as a forced lowering of temperature.

When a slug of hydroaromatic solvent and coal first experience heating in the preheater, the first reaction product is a gel which is formed in the temperature range 200 to 300C. Formation of the gel accounts for the first increase in Relative Viscosity". The gel forms due to bonding of the hydroaromatic compounds of the solvent with the hydrocarbonaceous material in the coal and is evidenced by a swelling of the coal. The bonding is probably a sharing of the solvent hydroaromatic hydrogen atoms between the solvent and the coal as an early stage in transfer of hydrogen from the solvent to the coal. The bonding is so tight that in the gel stage the solvent cannot be removed from the coal by distillation. Further heating of a slug in the preheater to 350C. causes the gel to decompose, evidencing completion of hydrogen transfer, producing a deashed solid fuel, liquid fuel and gaseous products and causing a decrease in Relative Viscosity".

A decrease of Relative Viscosity in the preheater is also caused by depolymerization of solvated coal polymers to produce free radicals therefrom. The depolymerization is caused by removal of sulfur and oxygen heteroatoms from hydrocarbonaceous coal polymers and by rupture of carbon-carbon bonds by hydrocracking to convert deashed solid fuel to liquid fuel and gases. The depolymerization is accompanied by the evolution of hydrogen sulfide, water, carbon dioxide, methane, propane, butane, and other hydrocarbons.

At the high temperatures of the preheater outlet zone, repolymerization of free radicals is a reaction which is favored over hydrogenation of free radical sites and accounts for the final tendency towards increase in "Relative Viscosity" in the preheater to a value above 10. This second increase in Relative Viscosity is avoided in accordance with the present invention. The elimination of sulfur and oxygen from the solvated deashed solid fuel is probably caused by stripping out of these materials by thermal rupture of bonds leaving free radical molecular fragments which have a tendency towards subsequent repolymerization at elevated temperature conditions. Termination of preheater conditions, such as a drop in stream temperature by forced cooling following the preheater step tends to inhibit polymer formation. The observed low level of sulfur in the liquid fuel product, which for one coal feed is about 0.3 weight percent, as compared to 0.7 weight percent in the vacuum bottoms (solid fuel) product, indicates that sulfur is being stripped out of the solid fuel product leaving low sulfur smaller molecular fragments as free radicals.

We have found that maximum or exit preheater temperatures should be in the range of 400 to 525C. The residence time in the preheater for a feed increment to achieve this maximum temperature is about 0.01 to 0.25 hours. At this combination of temperature and residence time, coke formation is not a problem unless flow is stopped, that is, unless the residence time is in creased beyond the stated duration. The hydrocarbon gas yield under these conditions is lessthan about 6 weight percent while excess solvent (liquid fuel) yield is above 10 or 15 weight percent, based on MAF coal feed, while the solid fuel product is above 20 weight percent. High production of gases is to be avoided because such production involves high consumption of hydrogen and because special facilities are required. However, a gaseous yield above 6 weight percent can be tolerated if facilities to store and transport the gas are available.

The relatively low sulfur content in the vacuum bottoms (deashed solid fuel) product of the present process is an indication that the reaction proceeds to a high degree of completion. It is also an indication that the vacuum bottoms product has been chemically released from the ash so that it can be filtered therefrom.

The hydrogen pressure in the present process is 35 to 300 kg/cm, generally, and 50 to 200 kg/cm' preferably. At about kg/cm hydrogen pressure, the solvent hydrogen content tends to adjust to about 6.1 weight percent. If the hydrogen content of the solvent is above this level, transfer of hydroaromatic hydrogen to the dissolved fuel tends to take place, increasing production of liquid fuel, which has a higher hydrogen content than solid fuel. If the solvent contains less than 6.1 weight percent of hydrogen, the solvent tends to acquire hydrogen from hydrogen gas at a faster rate than the fuel product. Once the solvent is roughly adjusted to a stable hydrogen level, conversion appears to depend on the catalytic effect of FeS, derived from the coal ash. Some deviations from this basic situation are observed in response to temperature and time variables. Higher temperatures tend to lower the hydroaromatic content of the system while rapid feed rates may preclude attainment of equilibrium values (not sufficient time). In addition, higher pressures tend to favor more rapid equilibrium and tend to increase the hydroaromatic character of the system.

In the dissolver stage of the present process, aromatic compounds which have surrendered hydrogen in the preheater are reacted with hydrogen to again form hydroaromatic compounds. Hydroaromatic compounds are partially (not completely) saturated aromatics. The chemical potential in the dissolver is too low for full saturation of aromatics to be a significant reaction. This is important because while hydroaromatics are capable of hydrogen transfer, saturated aromatics are not. Most of the saturates observed in the dissolver tend to be light products derived from ring opening of liquid product or solvent, or derived from aliphatic side chain removal. Solid fuel product aromatic species tend to remain aromatic or hydroaromatic.

The combined effect of time and temperature in the preheater stage is important in the present process. The desired temperature effect in the preheater stage is substantially a short time effect while the desired temperature effect in the dissolver requires a relatively longer residence time. The desired low preheater residence times are accomplished by utilizing an elongated tubular reactor having a high length to diameter ratio of at least 100, generally, and at least 1,000, preferably, so that rapidly upon reaching the desired maximum preheater temperature the preheater stream is discharged and the elevated temperature is terminated by forced cooling. Forced cooling can be accomplished by hydrogen quenching or by heat exchange. Thereupon, in the dissolver stage, wherein the temperature is lower, the

residence time is extended for a duration which is longer than the preheater residence time.

When the effluent of the present solvent refining coal process is distilled, normally gaseous materials are removed overhead. A middle boiling fraction comprising normally liquid carbonaceous fuel and (at start-of-run) solvent derived from a source outside of the process is also recovered. A residual high boiling material is also recovered which is solid at room temperature. The residual ash from the coal can be removed from the product prior to the product distillation by filtration, centrifuging or solvent extraction.

The middle boiling fraction of the vacuum distillation, which is liquid at room temperature, is normally at least partially recycled for slurrying with feed coal particles to function as the solvent for the process. Although even the vacuum bottoms, which is solid at room temperature, is inferior as a solvent to the normally liquid middle fraction of the vacuum distillation, it has been discovered that the ash is an advantageous recycle component. Apparently, the iron in the ash helps to break down extracted heavy coal molecules to provide an extracted fuel of lower average molecular weight. Therefore, when an ash-containing product stream is recycled, the relative yield of middle boiling fraction is enhanced while the relative yield of vacuum bottoms is decreased, as compared to the use of middle boiling fraction without ash containing recycle material as the solvent for the process. An additional advantage to the use of an ash-containing recycle stream is that subsequently produced vacuum bottoms has a significantly lower sulfur content.

Another advantage obtained by the use of an ashcontaining recycle stream is enhancement of activity for hydrogenation of the solvent to replenish hydrogen lost from the solvent by hydrogen donor activity. In order for hydrogen to react with the coal feed and accomplish solvent extraction of the coal from the ash residue and to accomplish removal of sulfur and oxygen heteroatoms from the extracted coal a transfer of hydrogen from partially hydrogenated aromatic solvent molecules is necessary. Gaseous hydrogen reacts indirectly with the coal and the extracted carbonaceous fuel from the coal by first chemically combining with the aromatic solvent molecules. An increased rate of combination of hydrogen with the aromatic solvent molecules in turn increases the rate at which the gaseous hydrogen reaches the coal. The initial hydrogenation activity of start-of-run solvent obtained from a source outside of the process is not of importance if sufficient time is allowed to permit the process to achieve a steady state by recycle of ash-containing process solvent. With continued recycle of ash-containing process solvent to the feed stream, the ash content in the feed slurry gradually builds up to an equilibrium level and as the equilibrium level is being achieved the benefits associated with the present invention become apparent.

The process of .the present invention utilizes a preheater, a dissolver and a vacuum distillation tower in series. A slurry of subdivided coal feed and recycle solvent is passed through the preheater. The preheater has a considerably smaller capacity than the dissolver so that the residence time in the preheater is considerably lower than the residence time in the dissolver. No external catalyst is added to either the preheater or dissolver stage, which are the only reactor stages of the present process. The effluent from the dissolver is filtered and then vacuum distilled. In the distillation column, gases are removed overhead, fuel product which is liquid at room temperature is removed from an intermediate position in the column, and fuel product which is solid at room temperature is removed from the bottom of the column; In accordance with the present invention, an advantageous modification of the process employs selective retention of mineral ash residue in the dissolver while permitting fuel product to preferentially leave the dissolver.

In the preheater the relative temperature is higher than in the dissolver while the residence time is lower than in the dissolver in order to enhance hydrogen transfer in the preheater from the recycle solvent to the feed coal, without excessive coking and polymerization. It has now been found that the minerals in the ash product, such as iron in the form of ferrous sulfide (FeS), catalyze the transfer of hydrogen from partially hydrogenated solvent produced in the dissolver to the feed coal in the preheater. As indicated above, the dissolver produces hydroaromatics rather than saturated aromatics since hydroaromatics are hydrogen donors while saturated compounds are not hydrogen donors and, therefore, their formation should be avoided. For this reason, in dissolver operation the use of a very efficient catalyst, a very high pressure or excessively long reaction times are self-defeating if they produce a high concentration of saturated compounds rather than hydroaromatics. In general, the dissolver employs a relatively lower temperature than the preheater and a re latively longer residence time by utilizing a larger reactor vessel than the preheater vessel.

It is advantageous to achieve a significant temperature differential between the preheater and the dissolver in order to increase the hydroaromatic activity in the process. In general, the advantages of a wide temperature differential are enhanced by an increase in the process of coal minerals, by a relatively high hydrogen partial pressure in the dissolver and a relatively lower hydrogen partial pressure in the preheater, by the use of dissolver temperatures as low as possible coupled with relatively longer dissolver residence times to compensate for low dissolver temperature and by the allowance of sufficient dissolver residence time to achieve a process equilibrium wherein the hydrogen transfer activity of the recycle solvent from the dissolver is considerably greater than hydrogen activity of the external solvent employed at process start-up.

The accumulation of minerals in the process not only enhances the chemical force of hydrogen transfer in the preheater but also assists the solvent in acquiring additional hydrogen in the next pass in the dissolver. Thereby, accumulation of coal minerals generally enhances conversion in the process. The present invention is directed towards retention of coal minerals in the dissolver stage without necessarily resorting to recycle.

The maximum temperature that can be employed in the preheater (which, because of plug flow, will occur at the preheater outlet) depends upon the hydrogen donor activity of the solvent. The relatively low temperatures employed in the dissolver relative to the preheater favor rehydrogenation of the solvent and favor the production of depolymerized coal-derived products in the dissolver. The maximum minerals concentration that could be tolerated in a minerals-containing recycle stream would depend on both the pumpability of the recycle stream and the filterability of the recycle stream because a portion of the recycle stream must be removed and filtered when the system achieves suitable process equilibrium. In accordance with the present invention, pumpability of the recycle stream is removed as a problem because according to the present invention the concentration of minerals is increased independently of minerals recycle.

Table l and FIG. I show that hydroaromatic activity of the solvent can be correlated with the iron content in the feed to the process. Various coals give data which fall on a smooth curve in this respect. As shown in FIG. 1, as minerals level increases there is an upward trend in solvent hydrogen level. The data in Table 1 were taken at 70 kg/cm hydrogen pressure and are illustrated in FIG. 2.

The data in Table l, as illustrated in FIG. 2, show that at a dissolver temperature of 450C. and a given iron level in the feed a lower hydroaromatic activity is imparted to the solvent as compared to a dissolver temperature of 425C, indicating that at any iron level in the feed a low dissolver temperature results in enhanced process hydroaromatic activity. it is desirable to maintain as high as possible a hydroaromatic activity in the solvent because the high hydroaromatic activity permits the use of a higher temperature in the preheater without coke and polymer formation becoming a problem. Since hydroaromatic activity is enhanced by operating the dissolver at relatively low temperatures for relatively longer holding times, it is important that the dissolver vessel be large compared to the preheater vessel. The use of differing temperatures in the preheater and in the dissolver, coupled with enhanced minerals concentration due to accumulation of ash in the dissolver, results in the preheater and the dissolver reactions occurring at preferred conditions for each. The separate types of reactions occurring in the preheater and the dissolver are further enhanced by employing a higher hydrogen pressure in the dissolver than in the preheater.

In the present process, oxygen is removed from the coal-derived fuel product more easily than sulfur, while nitrogen is bound in the most stable molecular structures. Therefore, as recycle of coal derived material increases it is found that nitrogen concentration in the stream increases while the sulfur concentration remains constant and the oxygen content decreases. Since high temperatures are the most favorable for sulfur removal, the higher the preheater temperature the lower will be the sulfur level in the product.

The first step in accomplishing solution of coal is the hydrogen transfer mechanism in the preheater. if this does not proceed adequately, the solution tends to degenerate by repolymerization and finally by coking. Many lignitic coals have only moderate amounts of iron, so that reactivity with hydrogen in the dissolver may not be satisfactory. If such coals contain high sodium levels, the use of carbon monoxide and steam to produce hydrogen in situ may be more satisfactory that the charging of hydrogen per se, since sodium is a catalyst for the shift reaction of carbon monoxide and steam to produce hydrogen. The rehydrogenation of the solvent in the dissolver may occur by using hydro gen with coal-derived ferrous sulfide as a catalyst or by using carbon monoxide and steam with sodium or ferrous sulfide as a catalyst.

in accordance with this invention, in an intermittent or continuous manner ash is selectively retained in the dissolver stage while liquid product is selectively permitted to leave. The amount of ash retained in the dissolver stage creates an ash concentration in the dis solver equivalent to the ash concentration in the dissolver that would be achieved if 10 to 80 percent of the ash in the feed coal, generally, were recycled, or if ID to 15 percent of the ash in the feed coal, preferably, were recycled. The retention of the ash in the dissolver in accordance with the present invention achieves the advantage of ash recycle but avoids the high pumping coats that would be required in order to pump an amount of ash slurry required to produce a similar concentration of ash in the dissolver vessel.

Table 2 shows the results of tests conducted with a split temperature process illustrating the effect of ash concentration when employing a preheater stage followed by a dissolver stage.

TABLE 2 TEST l 2 3 4 5 Ash Recycle No Yes Yes Yes Yes Pressure Hg, ltg/cm 70 70 70 70 Preheater Temp., "C. 450 450 450 450 475 Dissolver Temp, C. 425 425 425 425 425 l/LHSV', Hr. L79 1.89 l.72 1.79 1.79 GHSV 342 342 342 342 342 Solvent/MAP Coal/H,O 2.49/l/.05 2.49/l/.05 2.49/l/.05 2.49/1/00 2.49/1/00 Wt. Ash in Feed Slurry 5.285 7.42 9.25 10.08 10.65 Wt. Coal Derived Feed 34.8 48.7 60.7 66.l 69.9

YIELDS BASED ON MAF COAL FEED CO 0.5l 0.27 0.37 0.28

TABLE 2 Continued TEST 1 2 3 4 5 H,s 2.04 1.62 1.87 1.29 1.95 Hydrocarbon Gas 5.73 5.80 5.97 4.65 7.1 1 Gas Not Identified 59 9 30 1237 H1O 3.82 1.22 0.12 1.68 1.81 Excess Solvent (Liquid Product) 31.98 62.08 48.26 53.75 49.37 Vacuum Bottoms (Solid Product) 48.66 30.36 36.44 30.36 21.06 lnsol. Organic Matter 11.59 4.99 4.23 5.04 9.07 TOTAL 104.97 107.02 104.23 106.59 103.20

Recovery, weight 91: 95.91 92.63 92.73 96.48 93.35 MAF Conversion, weight 88.41 95.01 95.77 94.96 90.93

COMPOSITION OF LIQUID FUEL PRODUCT Carbon, weight 89.40 89.72 91.40 90.98 90.65 Hydrogen, weight 5.93 6.20 6.70 6.59 6.54 Nitrogen, weight 1.06 1.15 1.21 1.33 1.31 Sulfur, weight 0.410 0.420 0.358 0.338 0.438 Oxygen, weight 16 5.00 2.51 0.76 0.76 1.062

COMPOSITION OF VACUUM Bo'l'roMs FUEL PRODUCT Carbon, weight 88.54 88271 91.12 Hydrogen, wei ht k 4.74 5.35 5.10 Nitrogen, weig t b 2.22 2.10 2.22 Sulfur. weight 0.676 0.606 0.488 Oxygen, wei ht 3.619 3.156 1.00 Ash, weight 0.205 0.078 0.075

Test 1 of Table 2 was conducted with a solvent that liquid product tends to increase with increasing ash did not include recycled ash. in Tests 2, 3, 4 and 5 of content in the feed. As noted above, an increased hy- Table 2, the solvent included recycled ash. ln Test 1, drogen content in the liquid product signifies an inthe percent of ash in the feed slurry represents the ash creased heat content upon combustion. The data furfrom the feed coal, while in Tests 2, 3, 4 and 5 the perther show that with increasing ash content in the feed, cent ash in the feed slurry is continuously increasing the oxygen content of the filtrate product decreases. due to a progressively longer recycle duration. Al- Lower oxygen content in the fuel product also signifies though not yet achieved in the tests of Table 2, the proan increase in heat content. cess of Table 2 will ultimately achieve an equilibrium A highly important feature shown in Table 2 is that whereat the ash content in the feed slurry will stabilize the sulfur content of both the liquid product and the a d l v l ff, vacuum bottoms product decreases with increasing ash The yield data in Table 2 show that with increasing content in the feed slurry. Reduction in sulfur content ash content in the feed slurry the solvent is capable of of the vacuum bottoms product, which is solid deashed increased hydrogenation activity as evidenced by a coal at room temperature, imparts an enhanced comslight increase in hydrocarbon gas product, but more mercial value particularly to the solid deashed coal importantly by a great increase in production of liquid product. product coupled with a great decrease in solid fuel Table 2 shows that an optimum advantage in the re- (vacuum bottoms) product. The liquid product is a solcycle of ash-containing heavy vacuum bottoms is vent boiling range material which has a somewhat achieved by controlling the temperature in the pregreater hydrogen to carbon weight ratio and an 31 heater. Table 2 shows that gaseous hydrocarbon prodvated heat content (above 17,000 BTU/lb. or 9,450 uct decreases with increasing ash content in the feed cal./gm.), as compared to the solid fuel (vacuum bot- Slurry a preheater tempill'mufe of bul toms) product which has a somewhat lower hydrogen Crease5 at a preheater p of 475C to carbon ratio d a somewhat lower h content crease in hydrocarbon gas product is wasteful since it upon b ti b 16,000 BTU/|b or 8300 represents loss of desired liquid fuel product. A 6 cal./gm.). Although the recycle of ash results in slightly W g! P y l' f 8 y' an MAP basis more hydrogen consumption (2.5 to 3 percent), the hy- 1S Sullable pp limit gas Production unless gas drogen is gainfully utilized because of an increasing storage and transporting faclllties are available. Also, yield of higher heat content fuel. the llqllld product (excess solvent) yield diminishes Table 2 further shows that with increasing proporharply hen the preheater temperature increased tions of ash in the feed slurry, the percent of MAP cono to 475C. The percent MAF conversion also version tends to increase. The decrease in MAP conecreased when the preheater temperature was inversion in Test 5 is due to a combination of an excesc ased from 450 to 475C. sively high preheat temperature and an excessively lon Table 3 further illustrates the interdependence of the retention time in preheater. To avoid this drop in conrecycle 0f ash-containing vacuum bottoms product version, the preheat temperature should be below with the use of a higher temperature in the preheater 475C, preferably below 470 or 465C. and most prefthan in the dissolver. Tests 1, 2 and 3 of Table 3 did not erably no higher than 460 or 450C, or the retention employ recycle of ash while Tests 4 and 5 did. Table 3 time in the preheater should be decreased if higher temperatures are employed.

Table 2 also shows that the hydrogen content in the shows that Tests 4 and 5 exhibit progressively increasing ash contents in the feed slurry reflecting progressively extended ash-containing recycle durations.

TABLE 3 TEST 1 2 3 4 5 Ash Recycle No No No Yes Yes Pressure H ltg/cm 70 70 70 70 70 Preheater Temp., C. 450 500 450 450 475 Dissolver Temp., C. 450 450 425 425 425 l/LHSV; Hr. 0.52 0.98 1.79 1.89 1.79 GHSV 304 239 342 342 342 Solvent/MAP Coal/H 0 2.50] 110.08 2.49/1/006 2.49/1!0.05 2.49/ 11.05 2.49/1/00 Wt. Ash in Feed Slurry 5.0 5.0 5.285 7.42 10.65 Wt. Coal Derived Feed 33.3 33.3 34.8 48.7 69.9

M34515 Excess Solvent (Liquid Fuel Product) 5.36 15.10 31.98 62.08 49.37 Vacuum Bottoms (Solid Fuel Product) 68.12 56.81 48.66 30.36 21.06 lnsol. Organic Matter 14.91 13.83 1 1.59 4.99 9117 TOTAL 100.94 102.47 104.97 107.92 103.20

2L Recovery, weight 97.94 96.59 95.91 92.63 93.35 MAF Convemon, weight 85.09 86.17 88.41 95.01 90.93

COMPOSITION OF LIQUID FUEL PRODUCT Carbon, weight 89.68 89.40 89.72 90.65 Hydrogen, wei ht 5.94 5.93 6.20 6.54 Nltrogen, weig t 0.979 1.06 1.15 1.31 Sulfur, weight 0.46 0.410 0.420 0.438 Oxygen, weight 4.13 5.00 2.51 1.062

COMPOSITION OF SOLID FUEL PRODUCT Carbon, weight 87.32 89.03 88.54 88.71 91.12 Hydrogen, wei ht 5.11 5.12 4.74 5.35 5.10 Nitrogen, weig t 1.91 2.02 2.22 2.10 2.22 Sulfur, weight 0.944 0.719 0.676 0.606 0.488 Oxygen, weight it: 4.58 3.04 3.619 3.156 1.00 Ash, weight 0.133 0.067 0.205 0.078 0.075

Test 2 of Table 3 shows that employment of a split temperature between the preheater and the dissolver advantageously increases liquid product (excess solvent) yield, increases MAF conversion and decreases sulfur content in the vacuum bottoms product, as compared to Test 1 wherein the preheater and dissolver are operated at the same temperature. Moreover, Test 3 shows that when the preheater temperature and the dissolver temperature are split but the preheater temperature is not as high as 500C., the advantages tend to be enhanced even without solvent recycle. Tests 4 and 5 show that when ash recycle is employed with a split temperature the results are very much further improved in regard to excess solvent yield, MAF conversion and sulfur content in the vacuum bottoms, but the improvement in excess solvent yield and MAP conversion is sharply curtailed if the temperature in the preheater becomes excessive, even when vacuum bottoms ash-containing recycle is employed.

The tests of Table 3 indicate that best results are obtained employing ash recycle interdependently with a split temperature between the preheater and dissolver.

Table 4 shows the results of tests made to illustrate the effect of ash recycle upon the hydrogen content in the liquid product (excess solvent) of the process. Table 4 is an extension of the data of Tests 1 through 5 of Table 2 and shows that as the percent of iron present in the preheater feed increases due to continuing recycle of ash-containing vacuum bottoms, the hydrogen content in the solvent continuously increases above the hydrogen content of the original solvent employed TABLE 4 Fe in Increase in '7: H Test of Preheater Above H in Table 2 Feed Original Solvent FIG. 1 is a graphical illustration of the data of Table 4. The original solvent of the tests of Table 4 and FIG. 1 contained 6.04 weight percent hydrogen. As shown in Table 4 and FIG. 1 is a uniform increase in hydrogen in the liquid product corresponding to an increase in iron content in the feed slurry.

Table 2 and FIG. 2 both show that hydrogenation of the aromatic solvent in the dissolver is improved by operating the dissolver at a lower temperature than the preheater, thereby favoring the accumulation of high concentrations of hydroaromatic material in the dissolver product for recycle. The preheater is advantageously operated at a higher temperature than the dissolver to more efficiently transfer hydrogen from the hydrogenated aromatic to the coal feed in the next pass. In the preheater, high temperatures favor depolymerization, and favor sulfur and oxygen removal reactions. A high concentration of transferable hydrogen favors liquid formation and prevents coking. The maximum temperature allowed for the preheater depends on the activity of hydrogen in the solvent which is available for transfer. In the dissolver, moderate temperatures favor hydrogenation of both solvent and depolymerized coal. Also, in general, it is favorable to allow the catalytic ash minerals to build up to the highest concentration which can be managed with discharge pump and filter characteristics being the limiting factors. in accordance with the present invention, a savings in pumping costs is realized because ash minerals concentration is increased without resorting to ash recycle, or by only partially resorting to ash recycle.

FIG. 3 shows schematically the process of the present invention. As shown in FIG. 3, pulverized coal is charged to the process through line and contacted with recycle solvent from line 14 to form a slurring in contact with recycle hydrogen from line 40. The slurry passes through line 16 to preheater tube 18 having a high length to diameter ratio of at least 100 to permit plug flow. Preheater tube 18 is disposed in a furnace 20 so that in the preheater the temperature of a plug of feed slurry increases from a low inlet temperature to a maximum temperature at the preheater outlet.

The high temperature effluent slurry from the preheater is then passed through line 22 where it drops in temperature before reaching dissolver 24 due to the addition of cold makeup hydrogen through line 12. The residence time in dissolver 24 is substantially longer than the residence time in preheater 18 by virtue of the fact that the length to diameter ratio of dissolver 24 is considerably lower than preheater 18, causing backmixing and loss of plug flow. The slurry in dissolver 24 is at substantially a uniform temperature whereas the slurry in preheater l8 increeases in temperature from the inlet to the exit end thereof.

The data of FIG. 2 illustrate that the process is considerably improved by increasing the concentration of residue coal minerals in the dissolver zone. Although the concentration of coal minerals in the dissolver zone can be increased by solids recycle, this involves the costs and problems of solids slurry pumping. By employing cyclone separator 60 at the top of dissolver vessel 24, a portion of effluent solids can be separated from clear liquid product and directly retained in the dissolver vessel. In this manner, solids accumulation in dissolver 24 can be utilized to build up minerals or ash concentration to 2 to 20 weight percent, generally, or 5 to weight percent, preferably, without incurring unnecessary pumping costs. Furthermore, the presence of cyclone separator 60 permits variation and control of minerals concentration in dissolver 24 by utilizing partial or complete by-pass of the ashcontaining dissolver effluent around cyclone 60 through line 82.

Mineral residue and extracted liquid and solid fuel leaving dissolver 24 enter the circular truncated coni cal interior of cyclone separator 60 through tangential opening 62 in the lower wall thereof. The entering slurry swirls upwardly within the cyclone and deashed fuel is removed overhead through line 64 while solid is returned to dissolver 24 through line 66 having an adjustable value 68. A portion of the deashed fuel can be recycled as solvent, if desired, by passage through line 86 and valve 88. A portion of the effluent from dissolver 24 can by-pass cyclone 60, if desired, by passage through line 82 and valve 84. It is evident that the extent of solids retention in dissolver 24 can be controlled by manipulation of valves 68 and 84.

Solids or clear liquid leaving dissolver 24 through lines 64 and 82 pass through line 26 flash chamber 74. Unreacted hydrogen and gaseous hydrocarbons are removed overhead from flash chamber 74 through line '76 to distillation column 28. Ash-containing bottoms from flash chamber 74 is passed through line 78 to filter 80. Ash is removed from filter 80 through line 84 while filtrate is passed to distillation column 28 through line 82.

A distillate liquid product of the process is removed from a mid-region of distillation column 28 through line 42 and recovered as liquid product of the process. Since the process produces sufficient liquid to be withdrawn as liquid fuel product plus sufficient liquid to be recycled as solvent for the next pass, a portion of the liquid product is passed through line 44 for recycle to line 14, with or without fuel from line 86 to be employed to dissolve pulverized coal in the next pass.

From distillation column 28 gases, including hydrogen for recycle, are removed overhead through line 30 and are either withdrawn from the process through line 32 or passed through line 34 to scrubber 36 to remove impurities through line 38 and prepare a purified hydrogen stream for recycle to the next pass through line 40.

Vacuum bottoms is removed from distillation column 28 through line 46 and passed to conveyor belt 54 whereon the bottoms product is cooled to room temperature, at which temperature it solidifies. Solid, substantially ash-free fuel is removed from conveyor belt 54 by a suitable belt scrapper means, as indicated at 56. As shown in FIG. 7, no material is removed from the process between the preheater and the dissolver and all material entering the preheater passes through both the preheater and dissolver before any product separation occurs.

We claim:

1. A process for preparing deashed solid and liquid hydroocarbonaceous fuel from hydrocarbonaceous feed coal containing ash comprising contacting the feed coal with hydrogen and a solvent for the hydrocarbonaceous material in the coal to form a coal-solvent slurry in contact with hydrogen, passing the slurry and hydrogen through a preheater for a residence time between 0.01 and 0.25 hours, said preheater having a length to diameter ratio of at least to inhibit backmixing so that an increment of said slurry gradually increases in temperature in passage through the preheater from a low inlet temperature to a maximum temperature at the preheater outlet, the maximum temperature at the preheater outlet being 400 to 525C, the viscosity of an increment of the slurry in passage through the preheater increasing initially to a value at least 20 times the viscosity of the solvent alone when each is measured at a temperature of 99C., the viscosity of the slurry when measured at 99C. subsequently dropping to a value lower than 10 times the viscosity of the solvent alone when each is mesured at 99C. in continued passage through the preheater, the viscosity of said slurry finally tending to increase to a value greater than l times that of the solvent alone when each is measured at 99C. at the exit temperature of said preheater but the slurry and hydrogen being removed from said preheater after the relative viscosity drops to a value below but before the relative viscosity finally increases to a value of 10, passing the slurry to a dissolver maintained at a temperature between 350 and 475F. and which is below the temperature at the outlet of the preheater, the residence time of the slurry in the dissolver being greater than in the preheater, selectively retaining in the dissolver a portion of the ash removed from the coal relative to flow of fuel through the dissolver, removing dissolver effluent and separating said effluent into a gaseous stream, a fraction which is liquid at room temperature and a deashed fraction which is solid at room temperature, recycling hydrogen contained in said gaseous stream to the preheater, and recycling at least a portion of said liquid and/or solid fraction as solvent for said preheater step.

2. The process of claim 1 wherein selective ash retention results in an ash concentration of 2 to 20 weight percent in the dissolver.

3. The process of claim 1 wherein selective ash retention results in an ash concentration of 5 to weight percent in the dissolver.

4. The process of claim 1 including passing at ieast a portion of the dissolver effluent through a cyclone separator to achieve solids retention in the dissolver.

5. The process of claim 1 wherein the maximum temperature in the preheater is 425 to 500C.

6. The process of claim 1 wherein the temperature in the dissolver is 400 to 450C 7. The process of claim 1 including a forced temperature drop between the preheater and dissolver of at least C.

8. The process of claim 1 wherein the preheater length to diameter ratio is at least 1,000.

9. The process of claim 1 wherein the residence time in the preheater is 0.0l to 0.15 hours.

10. The process of claim I wherein the dissolver residence time is 0.1 to 3 hours.

1]. The process of claim 1 wherein carbon monoxide and steam are used together with or in place of hydrogen.

12. The process of claim 1 wherein the yield of deashed solid fuel is 20 to weight percent based on moisture and ash free coal feed.

13. The process of claim 1 wherein the viscosity of the slurry in the preheater falls to a value at least as low as 5 times the viscosity of the solvent alone when each is measureed at 99C.

14. The process of claim 1 wherein the viscosity of the slurry in the preheater falls to a value at least as low as 2 times the viscosity of the solvent alone when each is measured at 99C.

15. The process of claim 1 including a forced temperature reduction between the preheater and dissolver by injecting makeup hydrogen into the slurry stream between the preheater and dissolver.

16. The process of claim 1 including a forced temperature reduction between the preheater and dissolver by cooling the slurry stream in a heat exchanger between the preheater and dissolver.

17. The process of claim 1 wherein minerals concentration in the dissolver is equivalent to the minerals concentration that would be achieved by recycling 10 to 80 weight percent of the coal ash.

18. The process of claim 1 wherein production of hydrocarbon gaseous product comprises less than 6 weight percent based on moisture and ash free coal feed. 

1. A PROCESS FOR PREPARING DEASHED SOLID AND LIQUID HYDROCARBONACEOUS FUEL FROM HYDROCARBONACEOUS FEED COAL CONTAINING ASH COMPRISING CONTACTING THE FEED COAL WITH HYDROGN AND A SOLVENT FOR THE HYDROCARBONACEOUS MATERIAL IN THE COAL TO FORM A COAL-SOLVENT SLURRY IN CONTACT WITH HYDROGEN, PASSING THE SLURRY AND HYDROGEN THROUGH A PREHEATER FOR A RESIDENCE TIME BETWEEN 0.01 AND 0.25 HOURS, SAID PREHEATER HAVING A LENNGTH TO DIAMETER RATIO OF AT LEAST 100 TO INHIBIT BACKMIXING SO THAT AN INCREMENT OF SAID SLURRY GRADUALLY INCREASES IN TEMPERATURE IN PASSAGE THROUGH THE PREHEATER FROM A LOW INLET TEMPERATURE TO A MAXIMIUM TEMPERATURE AT THE PREHEATER OUTLET, THE MAXIMUM TEMPERATURE AT THE PREHEATER OUTLET BEING 400* TO 525*C., THE VISCOSITY OF AN INCREMENT OF THE SLURRY IN PASSAGE THROUGH THE PREHEATER INCREASING INITIALLY TO A VALUE AT LEAST 20 TIMES THE VISOSITY OF THE SOLVENT ALONE WHEN EACH IS MEASURED AT A TEMPERATURE OF 99*C., THE VISCOSITY OF THE SLURRY WHEN MEASURED AT 99*C. IN CONTINUED PASSAGE TO A VALUE LOWER THAN 10 TIMES THE VISCOSITY OF THE SOLVENT ALONE WHEN EACH IS MEASURED AT 99*C. IN CONTINUED PASSAGE THROUGH THE PREHEATER, THE VISCOSITY OF SAID SLURRY FINALLY TENDING TO INCREASE TO A VALUE GREATER THAN 10 TIMES THAT OF THE SOLVENT ALONE WHEN EACH IS MEASURED AT 99*C. AT THE EXIT TEMPERATURE OF SAID PREHEATER BUT THE SLURRY AND HYDROGEN BEING REMOVED FROM SAID PREHEATER AFTER THE RELATIVE VISCOSITY DROPS TO A VALUE BELOW 10 BUT BEFORE THE RELATIVE VISCOSITY FINALLY INCREASES TO A VALUE OF 10, PASSING THE SLURRY TO A DISSOLVER MAINTAINED AT A TEMPERATURE BETWEEN 350* AND 475*F. AND WHICH IS BELOW THE TEMPERATURE AT THE OUTLET OF THE PREHEATER, THE RESIDENCE TIME OF THE SLURRY IN THE DISSOLVER BEING GREATER THAN IN THE PREHEATER, SELECTIVELY RETAINING IN THE DISSOLVER A PORTION OF THE ASH REMOVED FROM THE COAL RELATIVE TO FLOW OF FUEL THROUGH THE DISSOLVER, REMOVING DISSOLVER EFFLUENT AND SEPARATING SAID EFFLUENT INTO A GASEOUS STREAM, A FRACTION WHICH IS LIQUID AT ROOM TEMPERATURE AND A DEASHED FRACTION WHICH IS SOLID AT ROOM TEMPERATURE, RECYCLING HYDROGEN CONTAINED IN SAID GASEOUS STREAM TO THE PREHEATER, AND RECYCLING AT LEAST A PORTION OF SAID LIQUID AND/OR SOLID FRACTION AS SOLVENT FOR SAID PREHEATER STEP.
 2. The process of claim 1 wherein selective ash retention results in an ash concentration of 2 to 20 weight percent in the dissolver.
 3. The process of claim 1 wherein selective ash retention results in an ash concentration of 5 to 15 weight percent in the dissolver.
 4. The process of claim 1 including passing at least a portion of the dissolver effluent through a cyclone separator to achieve solids retention in the dissolver.
 5. The process of claim 1 wherein the maximum temperature in the preheater is 425* to 500*C.
 6. The process of claim 1 wherein the temperature in the dissolver is 400* to 450*C.
 7. The process of claim 1 including a forced temperature drop between the preheater and dissolver of at least 20*C.
 8. The process of claim 1 wherein the preheater length to diameter ratio is at least 1,000.
 9. The process of claim 1 wherein the residence time in the preheateR is 0.01 to 0.15 hours.
 10. The process of claim 1 wherein the dissolver residence time is 0.1 to 3 hours.
 11. The process of claim 1 wherein carbon monoxide and steam are used together with or in place of hydrogen.
 12. The process of claim 1 wherein the yield of deashed solid fuel is 20 to 80 weight percent based on moisture and ash free coal feed.
 13. The process of claim 1 wherein the viscosity of the slurry in the preheater falls to a value at least as low as 5 times the viscosity of the solvent alone when each is measureed at 99*C.
 14. The process of claim 1 wherein the viscosity of the slurry in the preheater falls to a value at least as low as 2 times the viscosity of the solvent alone when each is measured at 99*C.
 15. The process of claim 1 including a forced temperature reduction between the preheater and dissolver by injecting makeup hydrogen into the slurry stream between the preheater and dissolver.
 16. The process of claim 1 including a forced temperature reduction between the preheater and dissolver by cooling the slurry stream in a heat exchanger between the preheater and dissolver.
 17. The process of claim 1 wherein minerals concentration in the dissolver is equivalent to the minerals concentration that would be achieved by recycling 10 to 80 weight percent of the coal ash.
 18. The process of claim 1 wherein production of hydrocarbon gaseous product comprises less than 6 weight percent based on moisture and ash free coal feed. 